Process for converting oxygenates to olefins with direct product quenching for heat recovery

ABSTRACT

The present invention relates to a process for converting an oxygenate in a feedstock to olefins in an olefin-containing effluent. Heat is transferred from effluent to the feedstock. The effluent is then quenched to form a light product fraction and a heavy product fraction. The heavy product fraction is fractionated to form a plurality of fractionated steams, at least one of which is used to heat the feedstock.

RELATED APPLICATIONS

This application is a divisional application of U.S. application Ser.No. 09/572,576, filed May 17, 2000, now U.S. Pat. No. 6,482,998; whichis a continuation-in-part of U.S. application Ser. No. 09/069,403, filedApr. 29, 1998, now U.S. Pat. No. 6,121,504.

FIELD OF THE INVENTION

The present invention relates to a process for increasing the efficiencyof heat recovery and improving heat integration with direct productquenching in the selective conversion of oxygenates to olefins.

BACKGROUND OF THE INVENTION

Light olefins (defined herein as ethylene, propylene, butenes andmixtures thereof) serve as feeds for the production of numerousimportant chemicals and polymers. Light olefins traditionally areproduced by cracking petroleum feeds. Because of the limited supply andescalating cost of petroleum feeds, the cost of producing olefins frompetroleum sources has increased steadily. Efforts to develop and improveolefin production technologies, particularly light olefins productiontechnologies, based on alternative feedstocks have increased.

An important type of alternative feedstocks for the production of lightolefins are oxygenates, such as alcohols, particularly methanol andethanol, dimethyl ether, methyl ethyl ether, methyl formate, anddimethyl carbonate. Alcohols may be produced by fermentation, or fromsynthesis gas derived from natural gas, petroleum liquids, carbonaceousmaterials including coal, recycled plastics, municipal wastes,agricultural products, or most organic materials. Because of the widevariety of raw material sources, alcohol, alcohol derivatives, and otheroxygenates have promise as an economical, non-petroleum feedstock sourcefor olefin production.

The conversion of oxygenates to olefins takes place at a relatively hightemperature, generally higher than about 250° C., preferably higher thanabout 300° C. Because the conversion reaction is exothermic, theeffluent typically has a higher temperature than the initial temperaturein the reactor. Many methods and/or process schemes have been proposedto manage the heat of reaction generated from the oxygenate conversionreaction inside of the reactor in order to avoid temperature surges andhot spots, and thereby to reduce the rate of catalyst deactivation andreduce the production of undesirable products, such as methane, ethane,carbon monoxide and carbonaceous deposits or coke. It would be veryuseful to have a process that effectively utilizes the heat of reactioncontained in the products exiting the oxygenate conversion reactor,optimizes heat recovery, and reduces overall utility consumption in theconversion of oxygenates to olefins. Such a process is environmentally,economically, and commercially more attractive.

In the conventional systems, the oxygenate conversion reaction ispredominantly conducted in the vapor phase using feedstocks and diluentsthat are usually liquid at ambient conditions. This requires supplyingsubstantial heat to the process to boil the oxygenate feedstock prior tointroducing it to the reactor, conventionally supplied by steam heatexchange or furnaces. Loss of energy is incurred in these indirect heatexchange methods, and substantial equipment is required. For steam,boilers must be built in addition to a steam/feed exchanger, andconstruction of a furnace is more expensive and complicated than atraditional heat exchanger. Methods are needed to improve the energyefficiency of the oxygenate conversion process and reduce the cost ofproviding vaporized oxygenate feedstock to an oxygenate conversionreactor.

Energy efficiency and cost of providing vaporized oxygenate feed isfurther complicated if utilization of a diluent is desired. The mostcommonly noted diluent, water/steam as disclosed in U.S. Pat. No.5,714,662, requires substantial energy and equipment cost to generate,but has the advantage of being easily able to separate from desiredlight olefins (especially ethylene and propylene). Other very commonlynoted diluents such as inert gases, including nitrogen, helium and evenmethane (see U.S. Pat. No. 5,744,680) require no energy or equipment tovaporize, but require extensive energy and equipment to separate fromthe desired light olefin product. Further, use of diluents can allowhigh total pressures while providing low oxygenate partial pressures,which can be advantageous in reducing compression energy needed in theoverall (including olefin separation and recovery) oxygenate conversionprocess, but this benefit may be outweighed by the energy costs ofboiling and separating of the diluent just noted. Proper selection ofdiluent composition is also needed to improve the energy efficiency inthe overall process and reduce the cost of providing vaporized oxygenatefeedstock to an oxygenate conversion reactor.

SUMMARY OF THE INVENTION

The present invention provides a process for converting an oxygenate toolefins with increased heat recovery and heat integration, said processcomprising: heating a feedstock comprising said oxygenate having a firstheat content from a first temperature to a second temperature throughfrom one to about three stages having successively higher heat contents;contacting said feedstock at said second temperature with a catalystcomprising a molecular sieve under conditions effective to produce adeactivated catalyst having carbonaceous deposits and a productcomprising said olefins, wherein said molecular sieve comprises poreshaving a pore diameter smaller than about 10 Angstroms and said producthas a third temperature which is higher than said second temperature;quenching said product with a medium at an initial temperature and in anamount sufficient for forming a light product fraction and a heavyproduct fraction wherein said light product fraction comprises lightolefins and said heavy product fraction has a final temperature which ishigher than said first temperature by at least about 5° C.; using saidheavy product fraction to provide heat at one or more of said stages toachieve said higher heat contents.

In another preferred embodiment, the process for converting an oxygenateto olefins comprises providing a feedstock comprising the oxygenate,transferring heat from at least a portion of an effluent of an oxygenateconversion reactor to the feedstock to cause at least a portion of thefeedstock to vaporize and form a vaporized feedstock, and contacting thevaporized feedstock at a temperature from about 200 to about 750° C. anda pressure from 1 kPa to 100 MPa with a catalyst comprising a molecularsieve having a pore diameter smaller than 10 Angstroms, wherein thefeedstock has a boiling range of no greater than about 30° C., theoxygenate conversion reactor converts at least a portion of thefeedstock into the effluent and the effluent comprises the olefins.

In another preferred embodiment, the process for converting an oxygenateto olefins comprises providing a feedstock comprising the oxygenate anda diluent, transferring heat from at least a portion of an effluent ofan oxygenate conversion reactor to the feedstock to cause at least aportion of the feedstock to vaporize and form a vaporized feedstock, andseparating the diluent from the effluent, wherein the feedstock has aboiling range of no greater than about 30° C., the oxygenate conversionreactor converts at least a portion of the feedstock into the effluent,and the effluent comprises the olefins.

In another preferred embodiment, the process for converting an oxygenateto olefins comprises providing a feedstock comprising the oxygenate anda diluent, transferring heat from at least a portion of an effluent ofan oxygenate conversion reactor to the feedstock to cause at least aportion of the feedstock to vaporize and form a vaporized feedstock,contacting the vaporized feedstock at a temperature from about 200 toabout 750° C. and a pressure from 1 kPa to 100 MPa with a catalystcomprising a molecular sieve having a pore diameter smaller than 10Angstroms, and separating the diluent from the effluent, wherein thefeedstock has a boiling range of no greater than about 30° C., theoxygenate conversion reactor converts at least a portion of thefeedstock into the effluent, and the effluent comprises the olefins.

BRIEF DESCRIPTION OF THE DRAWINGS

FIG. 1 is a flow diagram of a preferred embodiment of increasing heatrecovery in the present invention.

DETAILED DESCRIPTION OF THE INVENTION

The present invention provides a process for increasing heat recoveryand decreasing energy and utility requirements during the conversion ofoxygenates to olefins. The process involves taking the product mixture,including any unreacted oxygenate feed, from an oxygenate conversionreactor and, without fractionating the products, directly quenching theproduct mixture with a suitable medium, preferably water. This type ofquenching hereinafter will be referred to as “direct product quench.”The direct product quench removes heat from the product mixture, causinghigher boiling components, such as water and unreacted oxygenate feed,to condense and form a heavy product fraction. The heavy productfraction is separated from the light product fraction comprising gaseoushydrocarbon components such as light olefins, methane, ethane, propane,and butanes. The heavy product fraction may be divided further intoseveral fractions. The heavy production fraction, or any, or all of theseveral fractions may be subjected to various techniques or methods toseparate the quench medium from other components. The heavy productfraction, or any, or all of the several fractions or streams producedfrom quench medium separations thereof, may be used to supply at leastpart of the heat needed to vaporize or otherwise to increase the heatcontent of the oxygenate feedstock, through from one to about threestages, prior to being introduced into the oxygenate conversion reactorfor contacting the oxygenate conversion catalyst. These stages give theoxygenate feedstock successively higher heat content.

Most catalysts that are used in oxygenate conversion processes comprisemolecular sieves, both zeolitic (zeolites) and non-zeolitic types. Thepresent invention should achieve many of the desired improvements usingsubstantially any molecular sieve catalyst, regardless of the structuretype or pore size. Mixtures of zeolitic and non-zeolitic molecularsieves also may be used. Preferred molecular sieve catalysts for useaccording to the present invention comprise “small” and “medium” poremolecular sieve catalysts. “Small pore” molecular sieve catalysts aredefined as catalysts with pores having a pore diameter of less thanabout 5.0 Angstroms. “Medium pore” molecular sieve catalysts are definedas catalysts with pores having a pore diameter in the range of fromabout 5.0 to about 10.0 Angstroms. “Large pore” molecular sievecatalysts are catalysts with pores having a pore diameter larger thanabout 10.0 Angstroms. Generally, large pore molecular sieve catalysts,without additional appropriate modifications and/or treatments, are notpreferred catalysts for converting oxygenates to light olefins.

Zeolitic molecular sieve catalysts suitable for the use in the presentinvention with varying degree of effectiveness include, but are notnecessarily limited to AEI, AFI, CHA, ERI, LOV, RHO, THO, MFI, FER, andsubstituted examples of these structural types, as described in W. M.Meier and D. H. Olson, Atlas of Zeolitic Structural Types (ButterworthHeineman-third edition, 1997), incorporated herein by reference.Preferred zeolite catalysts include but are not necessarily limited tozeolite 3 A, zeolite 4 A, zeolite 5 A (collectively referred tohereinafter as zeolite A), ZK-5, ZSM-5, ZSM-34, erionite, chabazite,offretite, silicalite, borosilicates and mixtures thereof. See Meier andOlson. These zeolites may be obtained from many companies and commercialsources such as Mobil, AMOCO, UCI, Engelhard, Aldrich Chemical Company,Johnson Matthey Company, Union Carbide Corporation, and others.

Silicoaluminophosphates (“SAPO's”) are one group of non-zeoliticmolecular sieves that are useful in the present invention. SuitableSAPO's for use in the invention include, but are not necessarily limitedto SAPO-17, SAPO-18, SAPO-34, SAPO-44, and mixtures thereof Small poreSAPO's are preferred for producing light olefins. A preferred SAPO isSAPO-34, which may be synthesized according to U.S. Pat. No. 4,440,871,incorporated herein by reference, and Zeolites, Vol. 17, pp. 212-222(1996), incorporated herein by reference. SAPO-18 may be synthesizedaccording to J. Chen et al. Studies in Surface Sciences and Catalysis,Proceedings of the Tenth International Catalysis Society, Volume 84, pp17-31 (1994).

Substituted silicoaluminophosphates (substituted SAPO's) form anotherclass of non-zeolitic molecular sieves known as “MeAPSO's,” that aresuitable for use as catalysts in the present invention. MeAPSO's aredescribed in U.S. Pat. No. 4,567,029 and U.S. Pat. No. 5,126,308,incorporated herein by reference. SAPO's with substituents incorporatedafter synthesis, also may be suitable for use in the present invention.Suitable substituents, “Me,” include, but are not necessarily limited tonickel, cobalt, manganese, chromium, iron, zinc, strontium, magnesium,barium, and calcium. Preferred MeAPSO's include, but are not necessarilylimited to NiSAPO-17, NiSAPO-34, Co-SAPO-17, Co-SAPO-34, Sr modifiedSAPO-17 (SrAPSO-17), Sr modified SAPO-18 (SrAPSO-18), Sr modifiedSAPO-34 (SrAPSO-34), SrAPSO-44, and mixtures thereof. Differentsubstituents may be incorporated during or after the synthesis of thesilicoaluminophosphates.

Substituted aluminophosphates (ALPO) known as MeAPO's may also be usedas the non-zeolitic molecular sieve catalysts for the present invention.MeAPO's include, but are not necessarily limited to ZnAPO, ZrAPO, TiAPO,and mixtures thereof These molecular sieves may be synthesized accordingto U.S. Pat. Nos. 4,861,743, 4,567,029, 5,126,308.

Because the catalyst may be used in a variety of oxygenate conversionreactors and/or under a variety of reaction conditions, it may containbinders, fillers, or other material to provide better catalyticperformance, attrition resistance, regenerability, and other desiredproperties for a particular type reactor. When used in a fluidized bedreactor, the catalyst should be fluidizable under the reactionconditions. A catalyst may be subjected further to a variety oftreatments to achieve the desired physical, mechanical, and catalyticcharacteristics. Such treatments include, but are not necessarilylimited to calcination, milling, ball milling, grinding, spray drying,hydrothermal treatment with steam at elevated temperatures-from about400° C. to about 800° C., acid treatment, base treatment, andcombinations thereof.

The oxygenate feedstock of this invention comprises at least one organiccompound which contains at least one oxygen atom, such as aliphaticalcohols, ethers, carbonyl compounds (aldehydes, ketones, carboxylicacids, carbonates, esters and the like). When the oxygenate is analcohol, the alcohol can include an aliphatic moiety having from 1 to 10carbon atoms, more preferably from 1 to 4 carbon atoms. Representativealcohols include but are not necessarily limited to lower straight andbranched chain aliphatic alcohols and their unsaturated counterparts.Examples of suitable oxygenate compounds include, but are not limitedto: methanol; ethanol; n-propanol; isopropanol; C₄-C₂₀ alcohols; methylethyl ether; dimethyl ether; diethyl ether; di-isopropyl ether;formaldehyde; dimethyl carbonate; dimethyl ketone; acetic acid; andmixtures thereof. Preferred oxygenate compounds are methanol, dimethylether, or a mixture thereof.

The method of making the preferred olefin product in this invention caninclude the additional step of making these compositions fromhydrocarbons such as oil, coal, tar sand, shale, biomass and naturalgas. Methods for making the compositions are known in the art. Thesemethods include fermentation to alcohol or ether, making synthesis gas,then converting the synthesis gas to alcohol or ether. Synthesis gas canbe produced by known processes such as steam reforming, autothermalreforming and partial oxidization.

One or more inert diluents may be present in the feedstock, for example,in an amount of from 1 to 99 molar percent, based on the total number ofmoles of all feed and diluent components fed to the reaction zone (orcatalyst). As defined herein, diluents are compositions which areessentially non-reactive across a molecular sieve catalyst, andprimarily function to make the oxygenates in the feedstock lessconcentrated. Typical diluents include, but are not necessarily limitedto helium, argon, nitrogen, carbon monoxide, carbon dioxide, water,essentially non-reactive paraffins (especially the alkanes such asmethane, ethane, and propane), essentially non-reactive alkylenes,essentially non-reactive aromatic compounds, and mixtures thereof. Thepreferred diluents are water and nitrogen. Water can be injected ineither liquid or vapor form.

Hydrocarbons can also be included as part of the feedstock, i.e., asco-feed. As defined herein, hydrocarbons included with the feedstock arehydrocarbon compositions which are converted to another chemicalarrangement when contacted with molecular sieve catalyst. Thesehydrocarbons can include olefins, reactive paraffins, reactivealkylaromatics, reactive aromatics or mixtures thereof. Preferredhydrocarbon co-feeds include, propylene, butylene, pentylene, C₄⁺hydrocarbon mixtures, C₅ ⁺hydrocarbon mixtures, and mixtures thereof.More preferred as co-feeds are a C₄ ⁺hydrocarbon mixtures, with the mostpreferred being C₄ ⁺hydrocarbon mixtures which are obtained fromseparation and recycle of reactor product.

Preferably, the oxygenate feedstock should be at least partiallyvaporized and contacted in a suitable oxygenate conversion reactor withthe selected molecular sieve catalyst under process conditions effectiveto produce the desired olefins at an acceptable conversion level withdesired selectivities.

The temperature employed in the conversion process may vary over a widerange depending, at least in part, on the pressure, the selectedcatalyst, the reactor configuration, the weight hourly space velocity,and other reaction parameters. Although not limited to a particulartemperature, best results will be obtained if the process is conductedat temperatures in the range of from about 200° C. to about 750° C.,preferably in the range of from about 250° C. to about 650° C., and mostpreferably in the range of from about 300° C. to about 600° C.

Since the oxygenate feedstock normally is stored at ambient temperaturesbefore it is used in the conversion process, the feedstock has to beheated to a higher temperature with a much higher heat content suitablefor contacting the oxygenate conversion catalyst. It is preferable toincrease the heat content and/or the temperature of the feedstockthrough from one to about three intermediate stages, with each stagehaving a successively higher heat content. Many different streams in theoxygenate conversion process may be suitable sources for providing thenecessary heat to increase heat contents. These streams, including thosederived from the heavy product fraction from the quench tower and thestreams from the fractionator separating quench medium from othercomponents, are described in more detail below. It should be pointed outthat a stream may have a higher heat content after a heat exchange eventhough it has a lower temperature, primarily resulting from pressurechanges and/or phase changes, such as vaporization of a liquid. Thepressure and/or phase changes are needed for the oxygenate conversionprocess.

Light olefin products will form——although not necessarily in optimumamounts——at a wide range of pressures, including but not necessarilylimited to sub- and super-atmospheric pressures and autogeneouspressures, ——in the range of from about 1 kPa to about 100 MPa. Apreferred pressure is in the range of from about 5 kPa to about 50 MPa,most preferably in the range of from about 50 kPa to about 500 kPa. Theforegoing pressures are exclusive of diluent, if any is present, andrefer to the partial pressure of the feedstock as it relates tooxygenate compounds and/or mixtures thereof. Pressures outside of thestated ranges may be used and are not excluded from the scope of theinvention.

A steady state or semi-steady state production of light olefin productsmay be attained and/or sustained over a period of time, largelydetermined by the reactor type, the reactor configuration, thetemperature, the pressure, the catalyst selected, the amount of spentcatalyst recirculated (if any), the level of catalyst regeneration, theamount of carbonaceous materials left on the regenerated or partiallyregenerated catalyst, the weight hourly space velocity (WHSV), theamount of quench medium used, and other relevant process designcharacteristics.

A wide range of WHSV, defined as weight of total oxygenate feedstock perhour per weight of catalyst, for the feedstock will function in thepresent invention. Depending on the reactor type, the desired conversionlevel, the feedstock composition, and other reaction parameters, theWHSV generally should be in the range of from about 0.01 hr⁻¹ to about1000 hr⁻¹, preferably in the range of from about 0.1 hr⁻¹ to about 500hr⁻, and most preferably in the range of from about 0.5 hr⁻¹ to about200 hr⁻. Since the catalyst may contain other materials which act asinerts, fillers, or binders; the WHSV is calculated only on the weightbasis of oxygenate and molecular sieve part of the catalyst.

One or more diluents may be fed to the reaction zone with theoxygenates, such that the total feed mixture comprises diluent in arange of from about 1 mol % and about 99 mol %. Diluents which may beemployed in the process include, but are not necessarily limited to,helium, argon, nitrogen, carbon monoxide, carbon dioxide, hydrogen,water, paraffins, other saturated hydrocarbons (such as methane, ethane,propane, and mixtures thereof), aromatic compounds, and mixturesthereof. Preferred diluents include, but are not necessarily limited towater and nitrogen.

Oxygenate conversion should be maintained sufficiently high to avoid theneed for commercially unacceptable levels of recycling. 100% oxygenateconversion is preferred for the purpose of avoiding feedstock recyclecompletely. However, a reduction in unwanted by-products is observedfrequently when the oxygenate, particularly methanol, conversion levelis about 98% or less. Accordingly, there is usually from about 0.05 mol% to about 50 mol % unreacted oxygenate in the product stream along withthe oxygenate conversion products comprising olefins, water, and/orother byproducts. It is preferable to recover as much of the unreactedoxygenate as possible for recycle purposes. In any event, the oxygenatecontent in waste water may need to be reduced to an environmentallyacceptable level before byproduct water can be discharged.

Therefore, it is desirable to consider this incomplete oxygenateconversion in the overall heat recovery and heat integration scheme,i.e. optimizing heat recovery and heat integration, when using afractionator to recover unreacted oxygenates. If the oxygenateconversion level is high enough and/or recovery of unreacted oxygenateis not warranted for economic or environmental purposes, then thisinvention calls for utilizing heat directly from the heavy productfraction or any or all of the several fractions into which the heavyproduct fraction may be divided.

After contacting the oxygenate feed, the catalyst becomes fully orpartially deactivated due to accumulation of carbonaceous deposits (alsocalled coke) on the catalyst surface and/or inside the pores. Fullydeactivated catalyst The deactivated catalyst having carbonaceousdeposits is separated from the other oxygenate conversion products.Preferably at least a portion of the deactivated catalyst is separatedand withdrawn from the oxygenate conversion reactor intermittently,semi-continuously, continuously, or in batch. Before the deactivatedcatalyst is recycled back to the oxygenate conversion and used again, asuitable regeneration is carried out on at least a portion of thewithdrawn deactivated catalyst to remove at least a portion of thecarbonaceous deposits, in the range of from about 0.1 wt % to about 99.9wt %, preferably at least about 1.0 wt % of the carbonaceous depositsshould be removed. Complete regeneration——removing 100 wt % of theoriginal carbonaceous deposits on all of the deactivated catalyst—alsomay be carried out, but it is found that complete regeneration has atendency of leading to production of large amounts of undesirablebyproducts such as methane and/or hydrogen.

Preferably, the regeneration is carried out in the presence of a gascomprising oxygen or other oxidants. Air and air diluted with nitrogen,steam, and/or CO₂ are preferred regeneration gases. The catalystregeneration temperature should be in the range of from about 250° C. toabout 750° C., preferably from about 300° C. to about 700° C.

Almost any type of reactor will provide some conversions of theoxygenates to olefins. Reactor type includes, but is not necessarilylimited to fluidized bed reactor, riser reactor, moving bed reactor,fixed bed reactor, continuously stirred tank reactor, hybrids andcombinations thereof. Increased heat recovery and improved heatintegration in the present invention can be achieved with most anyreactor types. A preferred reactor system for the present invention is acirculating fluid bed reactor with continuous or semi-continuouscatalyst regeneration, similar to a modern fluid catalytic cracker.Fixed beds may be used, but are not preferred.

Because the oxygenate conversion reaction is highly exothermic, theoxygenate conversion reaction product effluent generally has a highertemperature than the feedstock temperature just before contacting thecatalyst. In one embodiment of the present invention, the feedstock fromthe storage tank at a first temperature and having a first heat contentis heated through several intermediate stages in heat exchangers to asecond desired temperature prior to contacting the oxygenate conversioncatalyst. It is preferable to have from one to about three stages ofheat exchange to provide streams with successively higher heat contents.Various streams from the oxygenate conversion process at differenttemperatures and external sources of heat, such as that from steam, maybe used as heat exchanger fluids to increase either the heat content,the temperature, or both, of the feedstock oxygenate.

After contacting the oxygenate feedstock with the oxygenate conversioncatalyst the oxygenate conversion reaction product effluent comprisingolefin products is quenched directly by contacting a suitable quenchmedium in a quench tower without first going through a productfractionation step. Alternatively, the product effluent may be used toprovide heat directly to the oxygenate feedstock. The temperature andthe heat content of the product effluent are reduced to intermediatelevels afterwards. The product effluent at this lower temperature andlower heat content is sent to the quench tower for direct quenching.

The compounds in the effluent stream which are gaseous under thequenching conditions are separated from the quench tower as a lightproduct fraction for olefin product recovery and purification. The lightproduct fraction comprises light olefins, dimethyl ether, methane, CO,CO₂, ethane, propane, and other minor components such as water andunreacted oxygenate feedstock. The compounds in the effluent streamwhich are liquid under quenching conditions, are separated from thequench tower as a heavy product fraction for heat recovery, and possibledivision into several fractions and separation of the quench medium. Theheavy product fraction comprises byproduct water, a portion of theunreacted oxygenate feedstock (except those oxygenates that are gasesunder quenching conditions), a small portion of the oxygenate conversionbyproducts, particularly heavy hydrocarbons (C5+), and usually the bulkof the quench medium.

Preferably, a quench medium is selected from a composition which remainssubstantially as a liquid under the quenching conditions, thusminimizing the amount of the quench medium present in the light gaseousproduct fraction which must undergo more expensive gaseous productprocessing steps to recover commercially acceptable grades of lightolefin products. A preferred quench medium is selected from the groupconsisting of water and streams that are substantially water. Morepreferably, the quench medium is a stream which is substantially waterand is selected from the several fractions of the heavy product fractionfrom the quench tower.

The amount of quench medium circulated in the quench tower at aparticular temperature for product quenching should be not more thanwhat is needed to produce a heavy product fraction exiting the quenchtower having a temperature at least about 5° C. higher than the firsttemperature of the oxygenate feedstock from the storage tank. In anotherembodiment, as already discussed, the oxygenate conversion reactoreffluent stream is used directly as a heat exchanger fluid to provideheat to the oxygenate feedstock before it enters the oxygenateconversion reactor to contact the oxygenate conversion catalyst.

Preferably, the pressure in the quench tower and the temperature of theheavy product fraction effluent are maintained at effective levels forrecovery of the highest quantity and quality of process heat. Morepreferably, the difference between the heavy product fraction effluentpressure and the pressure at which the feedstock is vaporized is belowabout 345 kPa, more preferably below about 207 kPa. The temperature ofthe heavy product fraction effluent from the quench tower preferably ismaintained at not less than about 30° C. below the bubble point of theheavy product fraction effluent. Maintaining a temperature differentialbetween the heavy product fraction effluent and its bubble pointprovides the highest possible bottoms temperature in the quench towerand the most economically practical recovery of useful heat from theheavy product fraction effluent.

Preferably, the heavy product fraction effluent (heavy product fraction)from the quench tower is pressurized and used for providing heat toother streams. In one embodiment, the heavy product fraction, or any, orall of the several fractions into which the heavy product fraction isdivided, or streams from quench medium separations thereof, are useddirectly as a heat exchanger fluid to increase the heat content and/ortemperature of the oxygenate feedstock at one or more of the stages withsuccessively higher heat contents. Further any of the several fractionsor streams produced from the quench medium separations thereof may beused to increase the heat contents of other streams within the overalloxygenate conversion reaction and product recovery process. The cooledquench medium recovered from such fractions and streams may be returnedback to the quench tower.

In a preferred embodiment, particularly when the oxygenate conversion isnot complete and the quench medium consists essentially of water, theheavy product fraction is divided into two fractions, a first fractionand a second fraction. The relative quantities of the first fraction andthe second fraction depend on the overall amount of heat that needs tobe removed from the product effluent stream in the quench operation, andthe temperature of the quench medium introduced into the quench tower.The relative quantities are set to optimize equipment cost for heatrecovery and utility consumptions. The first fraction is cooled to adesired temperature and sent back to the quench tower as a recycle, i.e.quench water. The utility required to cool the first fraction, e.g.cooling water, may be reduced by using the product effluent stream fromthe oxygenate conversion reactor as a heat exchange fluid to heat theoxygenate feedstock before the feedstock enters the oxygenate conversionreactor and/or before the product effluent stream enters the quenchtower.

The second fraction of the heavy product fraction effluent is sent to afractionator to separate the quench medium, which consists essentiallyof water——a part of it may originate as the recycled portion of thebyproduct water from the oxygenate conversion reaction when thefeedstock oxygenate has at least one oxygen——from other compounds, suchas unreacted oxygenates or certain heavier hydrocarbons from theoxygenate conversion reaction, present in the fraction. If other streamshaving compositions similar to or compatible with the second fractionexist within the oxygenate conversion and the associated productrecovery process, such other streams are combined with the secondfraction first and the combined stream is sent to the fractionator.

Generally it is desirable to fractionate a mixture into components assharply as possible. In the present invention, it is preferable for theoverhead oxygenate fraction and/or the heavies-containing fraction fromthe fractionator to have a composition of water as introduced in thesecond fraction of the heavy product fraction in the range of from about15 mol % to about 99.5 mol %, preferably from about 25 mol % to about 90mol %. An increase in the water composition of the overhead fractiontends to increase the condensation temperature, and more heat can berecovered economically from the overhead fraction of the fractionator toimprove heat integration for the entire process. Preferably, therecovered overhead oxygenate fraction contains at least about 90 mol %of the oxygenate contained in the second fraction of the heavy fraction.More preferably, the recovered overhead oxygenate fraction contains atleast about 99 mol % of the oxygenate contained in the second fractionof the heavy fraction.

The overhead fraction of the fractionator is condensed in a heatexchanger, i.e. a condenser, against the oxygenate feedstock at one ofthe stages, from one to about three where the oxygenate feedstock isgiven successively higher heat contents. Its preferable for the overheadfraction of the fractionator to have a pressure at least about 69 kPahigher than the pressure of the oxygen feedstock in the condenser. Thispressure differential also increases the condensation temperature of theoverhead fraction, making heat recovery from the overhead fraction moreeconomical.

The bottoms fraction of the fractionator consists essentially ofbyproduct water from the oxygenate conversion reaction. Preferably, thisbottoms fraction is pressurized and used to heat the oxygenate feedstockat one of the stages, from one to about three, where the oxygenatefeedstock is given successively higher heat contents prior to enteringthe oxygenate conversion reactor. The fractionator is operated such thatthe temperature of the bottoms fraction is at least about 5° C.,preferably at least about 25° C., higher than the first temperature ofthe oxygenate feed from storage. The operating temperature inside of thefractionator is determined by a number of parameters, including, but notnecessarily limited to the fractionator overhead pressure and theoverall pressure drop inside of the fractionator.

FIG. 1 shows one embodiment of a process flow diagram according to theinvention to increase heat recovery and to improve heat integration.Liquid oxygenate feed 1, such as methanol, having a first heat content,at a first temperature and a first pressure, is heated by stream 35 inheat exchanger 2. Stream 35 is fractionator bottoms stream 33 fromfractionator 24, which is pressurized by pump 34. The result is a firstheated oxygenate feed stream 3 with a higher heat content than that ofliquid oxygenate feed stream 1. First heated oxygenate feed stream 3then is heated in another heat exchanger 4 by overhead fraction 26 fromfractionator 24 to form a second heated oxygenate feed stream 5 with ahigher heat content than that of stream 3. Heat exchanger 4 is acondensor or a partial condenser for fractionator 24. Second heatedoxygenate feed stream 5 goes through steam pre-heater 6 to form a thirdheated oxygenate feed stream 7 which is further heated by oxygenateconversion product effluent 11 in heat exchanger 8 to form a fourthheated oxygenate feed stream 9 under the effectiveconditions——temperature, pressure, and proportion of liquid andvapor——desired for the conversion of the oxygenate feed. Oxygenateconversion product 11 is the effluent of oxygenate conversion reactor10, after being separated from the deactivated oxygenate conversioncatalyst which has carbonaceous deposits. Alternatively, heat exchanger8 may comprise of a plurality of coils inside of oxygenate conversionreactor 10.

Fourth heated oxygenate feed stream 9 is fed to oxygenate conversionreactor 10 which contains catalyst suitable for converting the oxygenatefeed to olefins. Oxygenate conversion reactor 10 may adopt variousconfigurations——fixed bed, fluidized bed, riser, moving bed, or acombination thereof, with or without continuous catalyst regeneration. Afixed bed reactor normally is not favored due to the difficulty ofwithdrawing deactivated catalyst for regeneration and returning theregenerated catalyst back to the reactor. The oxygenate feed isconverted to a product comprising light olefins and the catalyst becomesdeactivated or partially deactivated by accumulating carbonaceousdeposits which are formed as byproducts of the oxygenate conversionreaction.

Oxygenate conversion product effluent 11 flows through heat exchanger 8and becomes cooled oxygenate conversion product effluent stream 12 whichis sent to quench tower 13. Alternately, heat exchanger 8 may beeliminated and oxygenate conversion product effluent 11 is sent directlyto quench tower 13 without intermediate cooling. In quench tower 13oxygenate conversion product stream 12 contacts directly with a quenchmedium consisting essentially of water at an initial temperature over aseries of suitable contacting devices. The amount of the quench mediumneeded in quench tower 13 is dictated by a number of factors, including,but not necessarily limited to the composition of the quench medium, thetemperature of quench medium recycle introduced to quench tower 13, anddesired temperature differences and pressure differences between variousstreams. These differences are discussed where appropriate. The gaseousproducts are separated as light product fraction stream 14. Heavyproduct fraction stream 15, which exits from the bottom of the quenchtower at an exiting temperature, comprises the bulk of byproduct water,a portion of the unreacted oxygenate feedstock (except those oxygenatesthat are gaseous under the quenching conditions), a small portion of theoxygenate conversion byproducts, particularly heavy hydrocarbons (C5+),and usually the bulk of the quench medium.

A preferred quench medium is water, which is for all intents andpurposes indistinguishable from byproduct water. This eliminates theneed for steps to separate the quench medium from byproduct water in theheavy product fraction. In the event that a quench material other thanwater is used and this quench material is substantially in a liquid formunder quenching conditions, heavy product fraction 15, or any, or all ofthe several fraction into which the heavy product fraction is dividedmay be processed to separate the quench medium from byproduct water. Forexample, if the quench medium is a high boiling hydrocarbon such asdiesel fuel or similar streams, it is immiscible with byproduct water.Such a quench medium can be separated by a properly designed weir systemin the bottom of quench tower 13, or in an API separator or othersimilar devices at many different points of the process in the presentinvention. Further, if any heavy hydrocarbons (C5+) are formed in theoxygenate conversion reaction, they also may be removed from byproductwater in stream 15 or other points in the process in substantially thesame manner as or along with the removal of the quench medium. If thequench medium is a relatively light material which is substantiallygaseous under the quenching conditions, and hence being present insubstantial quantities in the light product fraction, such a quenchmedium can be separated in downstream olefin recovery processesencompassing the entire oxygenate conversion and olefin recovery andpurification process.

Regardless, the exiting pressure of heavy product fraction stream 15should be less than about 345×10³ pascals (345 kPa) below the pressureof liquid oxygenate feed i. Preferably, the exiting temperature of heavyproduct fraction stream 15 is maintained at not less than about 25° C.below the bubble point of byproduct water in stream 15. A preferredpressure difference between heavy product fraction stream 15 (lowerpressure) and liquid oxygenate feed 1 (higher pressure) is less than 207kPa.

Heavy product fraction stream (quench tower bottoms stream) 15 may beused to provide heat to the oxygenate feedstock in heat exchangers 2, 4,and/or 6 to increase the heat content of the feedstock. The oxygenatefeedstock contains successively higher heat contents at these stages.One or more of these stages also may be eliminated. Preferably, quenchtower bottoms stream 15 is divided into to two fractions, recyclefraction 18 and fractionator feed fraction 21. Recycle fraction 18, aquench water recycle stream, is cooled in exchanger 19 and recycled asquenching stream 20 back to quench tower 13. Alternatively, recyclefraction 18 or 20 may be split further into several fractions and thesefractions may be cooled to different temperatures in different heatexchangers. These fractions, or some of them, at different temperaturesmay be introduced into quench tower 13 at different points to betterintegrate heat recovery and minimize utility consumption. The heatcontent of fraction 18 may be used to provide heat to the oxygenatefeedstock in the heat exchanger 2, 4, and/or 6, or at differentlocations of the entire oxygenate conversion and olefin recovery andpurification process to provide heat and to increase heat recovery.

Fractionator feed fraction 21, optionally mixed with other watercontaining streams 22, is sent to fractionator 24. At least two streams,fractionator overhead stream 26 and fractionator bottoms stream 33, arefractionated from fractionator feed fraction 21. Fractionator overheadstream 26 should contain at least about 15 mol %, preferably at leastabout 25 mol %, of water from the oxygenate conversion reaction.Conjunctively with or alternatively to this composition preference, thetemperature of fractionator overhead stream 26 should be at least about10° C. higher than the boiling temperature of the oxygenate feed underthe conditions of heat exchanger 4.

Sufficient heat is added to fractionator 24 via reboiler 25, which whencoupled with a sufficient number of trays in fractionator 24 results inproducing fractionator bottoms stream 33 which comprises substantiallyall byproduct water and quench medium introduced with stream 23.

Preferably, the quench medium is water. When water is used as the quenchmedium, bottoms stream 33 consists essentially of the bulk of byproductwater from the oxygenate conversion reaction and no further steps arenecessary to separate byproduct water from the quench medium. If thequench medium is a material other than water and has not previously beenseparated from byproduct water prior to introduction into the quenchtower, this quench material may be separated from byproduct water inbottoms stream 33, or later in the process as described above. Further,if any heavy hydrocarbons (C5+) are formed in the oxygenate conversionprocess, they also may be removed from byproduct water in stream 33, orlater in the process in substantially the same manner as or along withthe removal of the quench medium.

Fractionator bottoms stream 33, before leaving fractionator 24, is at atemperature which is at least about 5° C., preferably at least about 25°C., higher than the first temperature of the oxygenate feed introducedfrom storage 1 to heat exchanger 2. The pressure at the top offractionator 24 should be at least 69 kPa higher than the pressure inheat exchanger 4 to increase heat recovery. Stream 35 is used to heat upliquid oxygenate feedstock 1 in heat exchanger 2. For better heatrecovery, exiting stream 36 from heat exchanger 2 preferably has atemperature equal to or less than about the temperature of stream 21.

One way to further improve heat integration and to increase heatrecovery is to use fractionator overhead stream 26 as the heat sourcefor heat exchanger 4. The cooled fractionator overhead stream 27 may befractionated further in separator 28 into vapor discharge stream 29 andliquid reflux 30 which is sent back to fractionator 24 after pressureadjustment with pump 31. It is important to maintain cooled fractionatoroverhead stream 27 at a temperature above the boiling point of the firstheated oxygenate feed 3 to provide favorable heat transfer.

Another embodiment of this invention relates to converting oxygenates toolefins with high energy and capital efficiency. In this preferredembodiment, an indirect heat transfer device for transferring heat fromat least a portion of an effluent of an oxygenate conversion reactor tothe feedstock is used to cause at least a portion of the feedstock tovaporize. For example, in FIG. 1, the indirect heat transfer devicewould be heat exchanger 8 and the effluent of the oxygenate conversionreactor is stream 10. In a preferred embodiment, the heat transferdevice is a thermosiphon utilizing a disengaging or circulation drum.

As defined herein, the boiling range of a feedstock is the difference inthe temperature of that feedstock at its dewpoint and its bubble pointat any one pressure of operation of the feedstock within the heattransfer device. If the feedstock is a single component feedstock, e.g.,a substantially pure methanol, the feedstock has a single temperatureboiling point at any one pressure. Thus it has a 0° C. boiling range.Typically, the feedstock will comprise one or more components havingdifferent boiling points, e.g., methanol, diluents, and otherhydrocarbon components. Typically, the feedstock will have more than onecomponent, and have a boiling point range of at least 2° C.

In a preferred embodiment, the boiling range is not greater than about30° C., 25° C., 20° C., 15° C., 10° C. and most preferably not greaterthan about 5 to 6° C. The oxygenate conversion reactor converts at leasta portion of the feedstock into the effluent, and the effluent comprisesthe olefins.

Desirably, the temperature of the reactor effluent is at least 300° C.,more preferably at least 350° C., and most preferably at least 400° C.The temperature of the reactor effluent preferably is below about 700°C. to achieve attractive yields in the oxygenate conversion reaction.

In another embodiment, a portion of the heat in the reactor effluentemanating from the indirect heat transfer device vaporizing thefeedstock is used to increase the sensible heat of the feedstock priorto vaporization of the feed stock. This heat is particularly useful foradding sensible heat (which is increasing the heat content to a liquidwithout vaporizing it, so the feedstock must be below its bubble point)to the feed stock.

In another embodiment, the sensible heat is provided in a separateindirect heat transfer device from that used for vaporization.Preferably, the sensible heat is added to the feedstock in a separateindirect heat transfer device prior to introducing it to thevaporization indirect heat transfer device.

In another embodiment, the reactor effluent provides heat ofvaporization or increase in sensible heat, to more than one feedstock inseparate indirect heat transfer devices. The indirect heat transferdevices that can be used to transfer heat include, for example, tubularexchangers, fin-type exchangers, condensers, scraped-surface exchangers,agitated vessels and thermosiphon-boilers. A thermosiphon-boiler is adevice wherein natural circulation of the boiling medium is obtained bymaintaining sufficient liquid head to provide for circulation, i.e.,circulation of feedstock through the device occurs by densitydifferences and is not forced by pumps. A separate vessel, desirably adrum in fluid communication with the exchanger, can be used to receivethe partially vaporized feedstock exiting the exchanger, separate theliquid and vapor, and return the liquid to the entrance of theexchanger. Fresh liquid or partially liquid feedstock can also be sentto the seperate vessel, rather than directly to the heat exchanger.

The thermosiphon-boiler is a particularly useful embodiment forvaporization of feedstock. Tubular exchangers include ashell-and-tube-type heat exchanger, a U-tube heat exchanger, apacked-lantern-ring exchanger, a outside-packed floating-head exchanger,an internal floating-head exchanger, a bent-tube fixed-tube-sheetexchanger, a bayonet-tube exchanger, a spiral-tube exchanger, afalling-film exchanger and Teflon-head exchanger.

The temperature of cooled reactor effluent is preferably not less than30° C. below the water dewpoint. In the method of this invention,determining the boiling range of a feedstock and the water dewpoint of acooled reactor effluent, and many other useful thermodynamic propertiesof the materials utilized, is desireably determined by the SoaveModified Redlich-Kwong (SMRK) equation of state. The SMRK equation isreadily available in SimSci PRO/II, which is a computer program forchemical process simulation.

In another embodiment, the reactor effluent stream can be split, usingone part to vaporize feedstock in one device. Other parts may be used toprovide heat to other materials in other devices. In a preferredembodiment, one part of the reactor effluent may be used to vaporizeoxygenate and another part may be used to vaporize diluent (or a mix ofdiluent and oxygenate) in another device, in a parallel approach,feeding both streams to the reactor. In another embodiment, all or partof the reactor effluent may be used to in one device to vaporizediluent, then the reactor effluent from that device is sent to anotherdevice to vaporize oxygenate (or a mix of diluent and oxygenate), or theorder of vaporization may be reversed, in a series approach. The bestapproach will be determined based upon the selection of oxygenate anddiluent employed, and other economic and process criteria, such asdesired energy efficiency and capital return requirements, usingengineering design and economic principles well known to those skilledin the art and not discussed in detail here.

In other embodiments, the diluent can have a normal boiling pointbetween −20° C. and 130° C.; preferably −6° C. and 100° C.; and mostpreferably 35° C. and 90° C. The diluent is desirably an aliphatic andaromatic hydrocarbons; C₄ to C₈ aliphatic and olefinic hydrocarbons andC₆-C₈ aromatics. Most preferably, the diluent is iso- or normal hexane.

It is particularly beneficial for the diluent to have a relatively lownormal boiling point, i.e., bubble point at one atmosphere pressure, toallow it to be vaporized at a temperature under the operating conditionsof the vaporizer that is low enough to provide high temperaturedifferentials with the reactor effluent. It should not, however, be solow that it cannot remain a liquid at the operating pressures ofvaporizer and reactor. Especially useful is when the diluent has anormal boiling point close to that of the oxygenate feed, so that it canbe mixed with the oxygenate in large proportion, if desired, and stillmaintain a low boiling range. This may make it possible to perform asmuch of the heating of the feedstock, both sensible heat andvaporization of both oxygenate and diluent, in as few exchangers aspossible.

The preferred diluents have desirable boiling ranges and normal boilingpoints. An added attractive feature is they have relatively low heats ofvaporization (compared to oxygenates and water/steam). This would allowthe heat from the reactor effluent to be most effectively utilized invaporizing the feedstock sent to the reactor, particularly that fractionwhich may be obtained in condensing the water out of the reactoreffluent, if so chosen. Iso- and normal hexane are particularly usefuldiluents when the oxygenate is methanol.

The process can be operated with a feedstock in which 0.1-100% of thefeedstock is in a liquid state. In other variations, the feedstock couldbe at least 10%, 20%, 30%, 40%, 50%, 60%, 70%, 80% and 90% in a liquidstate. Preferably, the temperature of the feedstock is not less than 30°C. lower than its bubble point at the location, and hence at thepressure, it is introduced to the heat exchange device in which it willbe vaporized, more preferably not less than 20° C. lower, still morepreferably not less than 10° C. lower, and most desirably no more than6° C. lower.

The invention will be better understood with reference to the followingexample, which illustrate, but should not be construed as limiting thepresent invention.

EXAMPLE I

A liquid methanol feed 1 at about 386.1 kPa pressure and 38° C. absorbsheat to increase its heat content in heat exchanger 2 from stream 35, at158° C. and 1,276 kPa pressure, from methanol/water fractionator 24 toform the first heated methanol feed stream 3 at a temperature of about100° C. and a pressure of 351.6 kPa. The first heated methanol feedstream 3 with 4,722 kJ/mole heat content absorbs heat from thefractionator overhead stream 26 in the heat exchanger 4 to form thesecond heated methanol feed stream 5 with a heat content of 6,521kJ/mole. Stream 5 is heated further by steam in heat exchanger 6 to formthe third heated methanol feed stream 7 which has even higher heatcontent than the third heated methanol feed stream 7-7,390 kJ/mole. Thethird heated methanol feed stream 7 is heated in heat exchanger 8 toform the fourth heated methanol feed stream 9 with the methanolconversion product effluent 11 from the oxygenate conversion reactor 10.The fourth heated methanol feed stream 9, having a much higher heatcontent of 17,102 kJ/mole, is suitable for contacting a catalyst in theoxygenate conversion reactor 10 to form a deactivated oxygenateconversion catalyst having carbonaceous deposits and a product 11comprising olefins, particularly light olefins. The oxygenate conversionreactor 10 is a fluidized bed reactor with continuous catalystregeneration and recirculation (not shown). The oxygenate conversionproduct 11 is separated from deactivated oxygenate conversion catalysthaving carbonaceous deposits and used to heat the stream 9 and form acooled methanol conversion product stream 12. A part of the deactivatedcatalyst is withdrawn and removed for regeneration. (not shown). It ispreferable to remove at least about 1.0 wt % of the carbonaceousdeposits from the deactivated catalyst during the regeneration. It isalso preferred to remove less than about 98.0 wt % of the carbonaceousdeposits from the deactivated catalyst during regeneration. Theregenerated catalyst is recycled back into the oxygenate conversionreactor 10 for contacting the oxygenate feed. 99.8 wt % of the methanolin stream 9 is converted in the reactor 10, with the unconverted balanceexiting in the stream 11.

The cooled methanol conversion product stream 12 exiting the heatexchanger 8 is sent to the quench tower 13, contacting directly a quenchmedium consisting essentially of water. The quench tower 13 is equippedwith suitable contacting devices inside. Most hydrocarbon products areseparated as a gaseous product stream 14. Heavier products, water, andunreacted methanol are discharge from the quench tower 13 as the quenchtower bottoms stream 15 at a temperature of about 116° C. and a pressureof about 262 kPa. The quench tower bottoms stream 15 is pressurized bythe pump 16 to form the pressurized quench tower bottoms stream 17 atabout 689.5 kpa. About 83 mol % of the pressurized quench tower bottomsstream 17 forms the recycle fraction 18 and is sent through the coolingexchanger 19 to form the quenching stream 20 at a lower temperature. Thequenching stream 20 is returned to the quench tower 13.

The rest of the pressurized quench tower bottoms stream 17, about 17 mol% becomes the fractionator feed fraction 21. The fractionator feedfraction 21 is combined with another methanol/water stream 22, a smallstream recovered from other sources within the overall oxygenateconversion and product recovery process. The combined stream 23 is sentto the fractionator 24. The fractionator overhead stream 26 containingabout 89 mol % water and about 10.5 mol % of methanol at a temperatureof 152° C. and a pressure at 551.6 kPa is sent to the heat exchanger 4.The bottoms from fractionator 24 is heated with steam in the heatexchanger 25 to produce the fractionator bottoms stream 33 at 158° C.and about 585.4 kPa, which contains primarily water with only traces ofother components. The fractionator bottoms stream 33 is pressurized toabout 1274.8 kPa and the resulting stream 35 is used for the heatexchanger 2 to heat the liquid methanol feed 1. After heat exchange, thebyproduct warm water stream 36 has a temperature of 46° C. at a pressureof 861.2 kPa.

Table 1 shows the product selectivity and the composition of productstream 11 of methanol conversion used for obtaining the results shown inTable 2 and Table 3. The feed rates, compositions, pressures, andtemperatures of various streams as described in Example I are shown inTable 2. The duties of key exchangers 2, 4, and 25 are tabulated inTable 3. Tables 2 and 3 were compiled using the Simulation Sciences,Inc. PRO/II chemical process simulation program utilizing the ModifiedPanagiotopoulos-Reid modifications to the Soave-edlich-Kwong equation ofstate.

TABLE 1 Product Selectivity Hydrocarbon Composition Component (wt %) inStream 11 (mol %) Hydrogen 0.15 0.73 Carbon Monoxide 0.03 0.01 CarbonDioxide 0.12 0.03 Methane 1.00 0.61 Ethylene 40.90 14.40 Ethane 0.830.27 Propylene 40.90 9.60 Propane 0.21 0.05 Butenes 8.89 1.56 Butanes0.09 0.02 Pentenes 3.95 0.56 Pentanes 0.04 0.01 Coke 2.89 — Total 100.0027.84

TABLE 2 Temper- Heat Stream Rate Methanol Water Pressure ature ContentNo. (mol/h) (mol %) (mol %) (kPa) (° C.) (kJ/mol)  1 10,000.0 98.23 1.77386.1 35.2 582  3 10,000.0 98.23 1.77 351.6 100.1 4,722  5 10,000.098.23 1.77 330.9 98.2 6,521  7 10,000.0 98.23 1.77 317.2 96.9 7,390  910,000.0 98.23 1.77 317.2 96.9 17,102 11 13,918.9 0.14 72.02 275.8 407.925,424 12 13,918.9 0.14 72.02 262.0 124.9 18,475 14 3,946.7 0.04 1.78241.3 37.8 6,841 15 92,909.4 0.18 99.82 262.0 115.6 3,856 21 9,972.20.18 99.82 689.5 115.6 3,859 22 863.8 0.21 99.78 689.5 43.4 1,394 261,118.0 10.54 89.42 551.6 151.9 21,740 27 1,118.0 10.54 89.42 517.1138.8 5,566 29 53.9 36.42 62.75 517.1 138.8 886 33 10,782.2 trace 100.00585.4 157.9 5,302 35 10,782.2 trace 100.00 1,274.8 158.1 5,310 3610,782.2 trace 100.00 861.2 46.1 1,476

TABLE 3 Exchanger No. Duty (10⁶ kJ/h) 2 41.4 4 18.9 6 8.6 8 97.1 19191.7 25 38.1

These results show that in the oxygenate conversion process, theexternal heat needed to bring the oxygenate feedstock to conditionsdesirable for contacting the catalyst, represented in the preferredembodiment by heat exchanger 6, is reduced as a result of increased heatrecovery and improved heat integration of the process.

Persons of ordinary skill in the art will recognize that manymodifications may be made to the present invention without departingfrom the spirit and scope of the present invention. The embodimentdescribed herein is meant to be illustrative only and should not betaken as limiting the invention, which is defined by the followingclaims.

We claim:
 1. A process for converting an oxygenate to olefina, theprocess comprising: providing a feedstock comprising the oxygenate,converting at least a portion of the oxygenate into an olefin-containingeffluent in an oxygenate conversion reactor, transferring heat from atleast a portion of the effluent to the feedstock to cause at least aportion of the feedstock to vaporize and form a vaporized feedstock,wherein the transferring heats the feedstock from a first temperature toa second temperature, quenching the effluent with a medium at an initialtemperature and in an amount sufficient for forming a light productfraction and a heavy product fraction wherein the light product fractioncomprises light olefins and the heavy product fraction has a finaltemperature which is higher than the first temperature by at least 5° C,fractionating at least a portion of the heavy product fraction in afractionator under conditions effective to form a plurality offractionated streams, and heating the feedstock with at least one of thefractionated streams.
 2. The process of claim 1, wherein the feedstockis selected from the group consisting of an impure methanol having aboiling range and a mixture of methanol and dimethyl ether.
 3. Theprocess of claim 1, wherein the oxygenate is selected from the groupconsisting of methanol, dimethyl ether, ethanol, methyl ethyl ether,dimethyl carbonate, methyl formate, methyl acetate, diethyl ether, andmixtures thereof.
 4. The process of claim 1, wherein the convertingoccurs in the presence of a catalyst selected from the group consistingof a zeolite, a silicoaluminophosphate (SAPO), a substitutedsilicoalumino-phosphate, a substituted aluminophosphate and mixturesthereof.
 5. The process of claim 4, wherein the catalyst comprises aSAPO selected from the group consisting of SAPO-17, SAPO-18, SAPO-34,SAPO-44, and mixtures thereof.
 6. The process of claim 1, wherein thefeedstock comprises a hydrocarbon diluent selected from the groupconsisting of C4 to C8 olefins, C4 to C8 aliphatics, C6-C8 aromatics andmixtures thereof.
 7. The process of claim 6, wherein the diluentcomprises a hydrocarbon selected from the group consisting of iso- andnormal hexane and mixtures thereof.
 8. The process of claim 1, whereinthe feedstock comprises a diluent having a normal boiling point of fromabout 35° C. to about 90° C.
 9. The process of claim 1, wherein thefeedstock comprises a boiling range of no greater than about 20° C. 10.The process of claim 9, wherein the boiling range is no greater thanabout 10° C.
 11. The process of claim 1, wherein the final temperatureis at least about 300° C.
 12. The process of claim 11, wherein the finaltemperature is at least about 350° C.
 13. The process of claim 1,wherein 0.1-100% of the feedstock is in a liquid state.
 14. The processof claim 13, wherein at least about 20% of the feedstock is in a liquidstate.
 15. The process of claim 14, wherein at least about 80% of thefeedstock is in a liquid state.
 16. The process of claim 1, wherein thetemperature of the feedstock is not lees than 30° C. lower than tobubble point of the feedstock at a location where the feedstock isintroduced to a heat transfer device that transfers heat from the atleast one of the fractionated streams to the feedstock.
 17. The processof claim 1, wherein at least a portion of the heat transferred from theat least one of the fractionated streams to the feedstock providessensible heat.
 18. The process of claim 1, wherein the temperature ofthe at least one of the fractionated streams after transferring heat tothe feedstock is not less than 30° C. below water dewpoint as determinedby the Soave Modified Redlich-Kwong equation of state.
 19. The processof claim 1, wherein the heat transferred from the at least one of tofractionated streams to the feedstock is done in an indirect heattransfer device.
 20. The process of claim 19, wherein the indirect heattransfer device is selected from the group consisting of ashell-and-tube-type exchanger and a thermosiphon-boiler.